Hydroprocessing of hydrocarbons

ABSTRACT

A process for hydrotreating (hydroprocessing) hydrocarbons and mixtures of hydrocarbons utilizing a catalytic composite of a porous carrier material, a Group VIII noble metal component and a germanium component, in which process there is effected a chemical consumption of hydrogen. A specific example of one such catalyst is a composite of a crystalline aluminosilicate, a platinum component and a germanium component, for utilization in a hydrocracking process. Other hydrocarbon hydroprocesses are directed toward the hydrogenation of aromatic nuclei, the ringopening of cyclic hydrocarbons, desulfurization, denitrification hydrogeneration, etc.

United States Patent 3,098,030 7/1963 Coonradtet al....

[54] HYDROPROCESSING OF HYDROCARBONS 9 Claims, No Drawings [52] U.S.Cl 208/111, 252/472, 252/461, 208/57, 208/89, 260/667 [51] 1nt.Cl ..C10g13/02 [50] Field of Search 208/] l 1;

[56] References Cited UNITED STATES PATENTS Primary ExaminerDelbert E. Gantz Assistant Examiner-R. M. Bruskin Attorneys-James R. Hoatson, Jr. and Robert W. Erickson ABSTRACT: A process for hydrotreating (hydroprocessing) hydrocarbons and mixtures of hydrocarbons utilizing a catalytic composite of a porous carrier material, a Group VIII noble metal component and a germanium component, in which process there is effected a chemical consumption of hydrogen. A specific example of one such catalyst is a composite of a crystalline aluminosilicatc, a platinum component and a germanium component, for utilization in a hydrocracking process. Other hydrocarbon hydroprocesses are directed toward the hydrogenation of aromatic nuclei, the ring-opening of cyclic hydrocarbons, desulfurization, denitrifcation hydrogeneration, etc.

HYDROPROCESSING OF HYDROCARBONS RELATED APPLICATION The present application is a continuation-in-part of my copending application, Ser. No. 828,762 filed May 28, 1969, all the teachings of which copending application are incorporated herein by specific reference thereto.

APPLICABILITY OF INVENTION The present invention encompasses the use of a catalytic composite of a porous carrier material, a Group VIII noble metal component and a germanium component in the hydrotreating of hydrocarbons and mixtures of hydrocarbons. As utilized herein, the term hydrotreating is intended to be synonymous with the term hydroprocessing", which involves the conversion of hydrocarbons at operating conditions selected to effect a chemical consumption of hydrogen. In cluded within the processes intended to be encompassed by the term hydroprocessing are hydrocracking, aromatic hydrogenation, a ring-opening, hydrorefining (for nitrogen removal and olefin saturation), desulfurization (often included in hydrorefining) and hydrogenation, etc. As will be recognized, one common attribute of these processes, and the reactions being effected therein, is that they are all hydrogen consuming", and are, therefore, exothermic in nature.

The individual characteristics of the foregoing hydrotreating processes, including preferred operating conditions and techniques, will be hereinafter described in greater detail. The subject of the present invention is the use of a catalytic composite which has exceptional activity and resistance to deactivation when employed in a hydrogen-consuming process. Such processes require a catalyst having both a hydrogenation function and a cracking function. More specifically, the present process uses a dual-function catalytic composite which enables substantial improvements in those hydroprocesses that have traditionally used a dual-function catalyst. The particular catalytic composite constitutes a porous carrier material, a Group VIII noble metal component and a germanium component; specifically, an improved hydrocracking process utilizes a crystalline aluminosilieate carrier material, a platinum component and a germanium component for improved activity, product selectivity and operational stability characteristics.

Composites having dual-function catalytic activity are widely employed in many industries for the purpose of accelerating a wide spectrum of hydrocarbon conversion reactions. Generally, the cracking function is thought to be associated with an acid-acting material of the porous, adsorptive refractory inorganic oxide type which is typically utilized as the carrier material for a metallic component from the metals, or compounds of metals, of Groups V through VIII of the Periodic Table, to which the hydrogenation function is generally attributed.

Catalytic composites are used to promote a wide variety of hydrocarbon conversion reactions such as hydrocracking, isomerization, dehydrogenation, hydrogenation, desulfurization, reforming, ring-opening, cyclization, aromatization, alkylation and transalkylation, polymerization, cracking, etc. some of which reactions are hydrogen producing while others are hydrogen consuming. In using the term hydrogen consuming, I intend to exclude those processes wherein the only hydrogen consumption involves the saturation of light olefins, resulting from undesirable cracking, which produces the light paraffins, methane, ethane and propane,. It is to the latter group of reactions, hydrogen consuming, that the present invention is applicable. In many instances, the commercial application of these catalysts is in processes where more than one of these reactions proceed simultaneously. An example of this type of process is a hydrocracking process wherein catalysts are utilized to effect selective hydrogenation and cracking of high molecular weight materials to produce a lower boiling, more valuable output stream. Another such example would be the conversion of aromatic hydrocarbons into LII jet fuel components, principally straight, or slightly branched paraffins.

Regardless of the reaction involved, or the particular process, it is very important that the catalyst exhibit not only the capability to perform its specified functions initially, but also perform them satisfactorily for prolonged periods of time. The analytical terms employed in the art to measure how efficient a particular catalyst performs its intended functions in a particular hydrocarbon conversion process, are activity, selectivity, and stability. For the purpose of discussion, these terms are conveniently defined herein, for a given charge stock, as follows: (1) activity is a measure of the ability of the catalyst to convert a hydrocarbon feed stock into products at a specified severity level, where severity level alludes to the operating conditions employedthe temperature, pressure,

I liquid hourly space velocity and hydrogen concentration; (2)

selectivity refers to the weight percent or volume percent of the reactants that are converted into the desired product and/or products; (3) stability connotes the rate of change of the activity and selectivity parameters with time-obviously, the smaller rate implying the more stable catalyst. With respect to a hydrogen-consuming process, for example hydrocracking, activity, stability and selectivity are similarly defined. Thus, activity connotes the quantity of charge stock, boiling above a given temperature, which is converted to hydrocarbons boiling below the given temperature. Selectivity refers to the quantity of converted charge stock which boils below the desired end point of the product, as well as above a minimum specified initial boiling point. Stability" connotes that rate of change of activity and selectivity. Thus, for example, where a gas oil, boiling above about 650 F., is subjected to hydrocracking, activity connotes the conversion of 650 F. plus charge stock to 650 F. minus product. Selectivity can allude to the quantity of conversion into gasoline-boiling range hydrocarbonsi.e., pentanes and heavier, normally liquid hydrocarbons boiling up to about 400 F. Stability might be conveniently expressed in terms of temperature increase required during various increments of catalyst life, in order to maintain the desired activity.

As is well known to those skilled in the art, the principal cause of observed deactivation or instability of a dual-function catalyst is associated with the fact that coke forms on the surface of the catalyst during the course of the reaction. More specifically, in the various hydrocarbon conversion processes, and especially those which are categorized as hydrogen consuming, the operating conditions utilized result in the formation of high molecular weight, black, solid or semisolid, hydrogen-poor carbonaceous material which coats the surface of the catalyst and reduces its activity by shielding its active sites from the reactants. Accordingly, a major problem facing workers in this area is the development of more active and selective catalytic composites that are not as sensitive to the presence of these carbonaceous materials and/or have the capability to suppress the rate of formation of these materials at the operating conditions employed in a particular process.

I have now found a dual-function catalytic composite which possesses improved activity selectivity and stability when employed in the hydroprocessing of hydrocarbons, wherein there is effected a chemical consumption of hydrogen. In particular, I have found that the use of a catalytic composite of a Group VIII noble metal component and a germanium component with a porous carrier material improves the overall operation of these hydrogen-consuming processes. Moreover. I have determined that a catalytic composite of a crystalline aluminosilicate carrier material, a platinum component and a germanium component, when utilized in a process for hydrocracking hydrocarbonaceous material into lower boiling hydrocarbon products, affords substantial improvement in performance and results. As indicated, the present invention essentially involves the use of a catalyst in which a germanium component has been added to a dual-function conversion catalyst, and enables the performance characteristics of the process to be sharply and materially improved.

An essential condition associated with the acquisition of this improved performance is the oxidation state of the germanium component utilized in this catalyst. As a result of my investigations, l have determined that the germanium component must be utilized in a positive oxidation state (i.e., either 2 or :ir4) and that the germanium component must be uniformly distributed throughout the porous carrier material. Furthermore, the catalyst must be prepared under carefully controlled conditions as will be explained hereinafter. In the case of a hydrocracking process, one of the principal advantages as sociated with the use of the novel catalyst of the present invention involves the acquisition of the capability to operate in a stable manner in a high-severity operation. in short, the present invention essentially involves the finding that the addition of a controlled amount of a germanium component, in a positive oxidation state, to a dual-function hydrocarbon conversion catalyst containing a Group VIII noble metal component enables performance characteristics of the catalyst to be sharply and materially improved when used in a hydrogenconsuming process.

OBJECTS AND EMBODIMENTS An object of the present invention is to afford a process for the hydroprocessing of a hydrocarbon, or mixtures of hydrocarbons. A corollary objective is to improve the selectivity and stability of hydroprocessing utilizing a highly active, germanium component-containing catalytic composite.

A specific object of my invention resides in the improvement of hydrogen-consuming processes including hydrocracking, hydrorefining, ring opening for jet fuel production, hydrogenation of aromatic hydrocarbons, desulfurization, denitrification, etc. Therefore, in one embodiment, the present invention encompasses a hydrocarbon hydroprocess which comprises reacting a hydrocarbon with hydrogen at conditions selected to effect chemical consumption of hydrogen and in contact with a catalytic composite of a Group VIII noble metal component, a germanium component and a porous carrier material. In another embodiment, the operating conditions include a pressure of from 400 to about 5,000 p.s.i.g., an LHSV (defined as volumes of liquid hydrocarbon charge per hour per volume of catalyst disposed in the reaction zone) of from 0.1 to about 10.0, a hydrogen concentration of from 1,000 to about 50,000 s.c.f./Bbl. and a maximum catalyst temperature of from 200 F. to about 900 F In another embodiment, the process is further characterized in that the catalytic composite is reduced and sulfided prior to contacting the hydrocarbon feed stream. ln still another embodiment, my invention involves a process for hydrogenating a coke-forming hydrocarbon distillate containing di-olefinic and mono-olefinic hydrocarbons, and aromatics, which process comprises reacting said distillate with hydrogen, at a temperature below about 500 F., in contact with a catalytic composite of an alumina containing refractory inorganic oxide, a Group VIII noble metal component, an alkali metal component and a germanium component, and recovering an aromatic/mono-olefinic hydrocarbon concentrate substantially free from conjugated di-olefinic hydrocarbons.

Another embodiment affords a catalytic composite comprising a substantially pure crystalline aluminosilicate material, at least about 90.0 percent by weight of which is zeolitic, a Group Vlll noble metal component and a germanium component.

Other objects and embodiments of my invention relate to additional details regarding preferred catalytic ingredients, the concentration of components in the catalytic composite, methods of catalyst preparation, individual operating conditions for use in the various hydrotreating processes, preferred processing techniques and the like particulars which are hereinafter given in the following, more detailed summary of my invention.

SUMMARY OF INVENTION As hereinabove set forth, the present invention involves the hydroprocessing of hydrocarbons and mixtures of hydrocarbons, utilizing a particular catalytic composite. This catalyst comprises a porous carrier material having combined therewith a GROUP Vlll noble metal component and a germanium component; in many applications, the catalytic will also contain a halogen component and in come select applications, an alkali metal or alkaline-earth metal component. Considering first the porous carrier ma material it is preferred that it be a porous, adsorptive, high surface area support having a surface area of about 25 to about 500 square meters per gram. The porous carrier material is necessarily relatively refractory with respect to the operating conditions employed in the particular hydrotreating process, and it is intended to include carrier materials which have traditionally been utilized in dualfunction hydrocarbon conversion catalysts. in particular, suitable carrier materials are selected from the group of amorphous refractory inorganic oxides including alumina, titania, zirconia, chromia, magnesia, thoria, boria, silica-alumina, silica-magnesia, chromia-alumina, alumina-boria, alumina-silica-boron phosphate, silica-zirconia, etc. When of the amorphous type, the preferred carrier material is a composite of alumina and silica with silica being present in an amount of about 10.0 percent to about 90.0 percent by weight.

In many hydroprocessing applications of the present invention, particularly hydrocracking heavy hydrocarbonaceous material to produce lower boiling hydrocarbon products, the carrier material will constitute a crystalline aluminosilicate, often referred to as being zeolitic in nature. This may be naturally occurring, or synthetically prepared, and includes mordenite, faujasite, Type A or Type U molecular sieves, etc. When utilized as the carrier material, the zeolitic material may be in the hydrogen form, or in a form which has been treated with multivalent cations.

As hereinabove set forth, the porous carrier material, for use in the process of the present invention, is a refractory inorganic oxide, either alumina in and of itself, or in combination with one or more other refractory inorganic oxides, and particularly in combination with silica. When utilized as the sole component of the carrier material, the alumina may be of the gamma-, eta-, or theta-alumina type, with gamma-, or eta-alumina giving the best results. ln addition, the preferred carrier materials have an apparent bulk density of about 0.30 to about 0.70 gm./cc. and surface area characteristics such that the average pore diameter is about 20 to about 300 Angstroms, the pore volume is about 0.10 to about 1.0 milliliters per gram and the surface area is about to about 500 square meters per gram. Whatever type of refractory inorganic oxide is employed, it may be activated prior to use by one or more treatments including drying, calcination, steaming, etc. For example, the alumina carrier may be prepared by adding a suitable alkaline reagent, such as ammonium hydroxide, to a salt of aluminum, such as aluminum chloride, aluminum nitrate, etc., in an amount to form an aluminum hydroxide gel which, upon drying and calcination, is converted to alumina. The carrier material may be formed in any desired shape such as spheres, pills, cakes, extrudates, powders, granules, etc, and may further be utilized in any desired size.

When a crystalline aluminosilicate, or zeolitic material, is intended for use as the carrier, it may be prepared in a number of ways. One common way is to mix solutions of sodium silicate, or colloidal silica, and sodium aluminate, and allow these solutions to react to form a solid crystalline aluminosilicate. Another method is to contact a solid inorganic oxide, from the group of silica, alumina, and mixtures thereof, with an aqueous treating solution containing alkali metal cations (preferably sodium) and anions selected from the group of hydroxyl, silicate and aluminate, and allow the solid-liquid mixture to react until the desired crystalline aluminosilicate has been formed. One particular method is especially preferred when the carrier material is intended to a crystalline aluminosilicate. This stems from the fact that the method can produce a carrier material of substantially pure crystalline aluminosilicate particles. In employing the term substantially pure", the intended connotation is an aggregate particle at least 90.0 percent by weight of which is zeolitic. Thus this carrier is distinguished from an amorphous carrier material, or prior art pills and/or extrudates in which the zeolitic material might be dispersed within an amorphous matrix with the result that only about 40.0 percent to about 70.0 percent by weight of the final particle is zeolitic. The preferred method of preparing the carrier material produces crystalline aluminosilicates of the faujasite modification, and utilizes aqueous solutions of colloidal silica and sodium aluminate. Colloidal silica is a suspension in which the suspended particles are present in very finely divided formi.e. having a particle size from about I to about 500 millimicrons in diameter. The type of crystalline aluminosilicate which is produced is primarily dependent upon the conditions under which crystallization occurs, with the Sim/A1 ratio, the Na 0/Si0 ratio, the H 0/Na 0 ratio, temperature and time being the important variables.

After the solid crystalline aluminosilicate has been formed, the mother liquor is separated from the solids by methods such as decantation or filtration. The solids are water washed and filtered to remove undesirable ions, and to reduce the quantity of amorphous material, and are then reslurried in water to a solids concentration of about 5.0 percent to about 50.0 percent. The cake and the water are violently agitated and homogenized until the agglomerates are broken and the solids are uniformly dispersed in what appears to be a colloidal suspension. The suspension is then spray dried by conventional means such as pressuring the suspension through an orifice into a hot, dry chamber. The solid particles are withdrawn from the drying chamber and are suitable for forming into finished particles of desired size and shape. The preferred form of the finished particle is a cylindrical pill, and these may be prepared by introducing the spray-dried particles directly into a pilling machine without the addition of any extraneous lubricant or binder. The pilling machines are adjusted to produce particles having a crushing strength of from 2 to 20 pounds, and preferably from 5 to pounds. The pilled faujasite carrier material, of which at least about 90.0 percent by weight is zeolitic, is activated catalytically by converting the sodium form either to the divalent ion form, the hydrogen form or mixtures thereof.

One essential constituent of the composite of the present invention is a germanium component, and it is an essential feature of the catalyst used in hydroprocessing according to the present invention, that the germanium component is present in the composite in an oxidation state above that of the elemental metal. That is to say, the germanium component necessarily exists within the catalytic composite in either the +2or +4oxidation state, the latter being the most likely state. Accordingly, the germanium component will be present in the composite as a chemical compound, such as the oxide, sulfide, halide, etc., wherein the germanium is in the required oxidation state, or as a chemical combination with the carrier material in which combination the germanium exists in this higher oxidation state. On the basis of the evidence currently available, it is believed that the germanium component in the subject composite exists as germanous or germanic oxide. It is important to note that this limitation on the state of the germanium component requires extreme care in the preparation and use of the subject composite in order to insure that it is not subjected to high-temperature reduction conditions (reduction at temperatures above l,000 F.) effective to produce the germanium metal. This germanium component may be incorporated in the catalytic composite in any suitable manner known to the art such as by coprecipitation or cogellation with the porous carrier material, ion exchange with the gelled carrier material or impregnation with the carrier material either after or before it is dried and calcined. It is to be noted that it is intended to include within the scope of the present invention all conventional methods for incorporating a metallic component in a catalytic composite and the particular method of incorporation used is not deemed to be an essential feature of the present invention. One method of incorporating the germanium component into the catalytic composite involves coprecipitating the germanium component during the preparation of the carrier material. This method typically involves the addition ofa suitable soluble germanium compound such as germanium tetrachloride to the inorganic oxide hydrosol and then combining the hydrosol with a suitable gelling agent and dropping the resultant mixture into an oil bath maintained at elevated temperatures. The droplets remain in the oil bath until they set and form hydrogel spheres. The spheres are withdrawn from the oil bath and subjected to specific aging treatments in oil and in an ammoniacal solution. The aged spheres are washed and dried at a temperature of about 200 F. to 400 F., and thereafter calcined at an elevated temperature of about 850 F. to about 1,300 F. Further details of spherical particle production may be found in U.S. Pat. No. 2,620,314, issued to James Hoekstra. After drying and calcining the resulting gelled carrier material there is obtained an intimate combination of alumina and germanium oxide. A preferred method of incorporating the germanium component into the catalytic composite involves utilization of a soluble, decomposable compound of germanium to impregnate the porous carrier material. In general, the solvent used in this impregnation step is selected on the basis of the capability to dissolve the desired germanium compound and is preferably an aqueous, or alcoholic solution. Thus, the germanium component may be added to the carrier material by commingling the latter with a solution of a suitable germanium salt or suitable compound of germanium such as germanium tetrachloride, germanium difluoride, germanium tetrafluoride, germanium di-iodide, germanium monosulfide, and the like compounds. In general, the germanium component can be impregnated either prior to, simultaneously with, or after the Group Vlll noble metal component is added to the carrier material. However, I have found that excellent results are obtained when the germanium component is impregnated simultaneously with the group Vlll noble metal component. In fact, I have determined that a preferred impregnation solution contains chloroplatinic acid, hydrogen chloride, and germanous oxide dissolved in chlorine water, especially when the catalyst is intended to contain combined chlorine. Following the impregnation step, the resulting composite is dried and calcined as explained hereinafter.

Regardless of which germanium compound is used in the preferred impregnation step, it is important that the germanium component be uniformly distributed throughout the carrier material. It is preferred to use a volume ration of impregnation solution to carrier material of at least l.5:l and preferably about 2:1 to about l0:l, or more. Similarly, it is preferred to use a relatively long contact time during the impregnation step ranging from about A hour up to about 7.; hour or more before drying to remove excess solvent in order to insure a high dispersion of the germanium component on the carrier material. The carrier material is, likewise, preferably constantly agitated during this preferred impregnation step.

As previously indicated, the catalyst for use in the process of the present invention also contains a Group Vlll noble metal component. Although the process of the present invention is specifically directed to the use of a catalytic composite containing platinum, it is intended to include other Group Vlll noble metals such as palladium, rhodium, ruthenium, osmium and iridium. The Group Vlll noble metal component, for example platinum, may exist within the final catalytic composite as a compound such as an oxide, sulfide, halide, etc., or in an elemental state. The Group Vlll noble metal component generally comprises about 0.01 percent to about 2.0 percent by weight of the final composite, calculated on an elemental basis. Excellent results are obtained when the catalyst contains about 0.3 percent to about 0.9 percent by weight of the Group Vlll noble metal. In addition to platinum, another particularly preferred Group Vlll noble metal component is palladium, or a compound of palladium.

The Group Vlll noble metal component may be incorporated within the catalytic composite in any suitable manner including coprecipitation or cogellation with the carrier material, ion exchange, or impregnation. A preferred method of preparation involves the utilization of a water-soluble compound of a Group Vlll noble metal component in an impregnation technique. Thus, a platinum component may be added to the carrier material by commingling the latter with an aqueous solution of chloroplatinic acid. Other water-soluble compounds of platinum may be employed, and include ammonium chloroplatinate, platinum chloride, dinitro diamino platinum, etc. The use of a platinum chloride compound, such as chloroplatinic acid, is preferred since it facilitates the incorporation of both the platinum component and at least a minor quantity of the halogen component in a single step. In addition, it is generally preferred to impregnate the carrier material after it has been calcined in order to minimize the risk of washing away the valuable Group Vlll noble metal compounds; however, in some instances it may prove advantageous to impregnate the carrier material when it exists in a gelled state. Following impregnation, the composite will generally be dried at a temperature of about 200 F. to about 400 F., for a period of from 2 about 24 hours, or more, and finally calcined at a temperature of about 700 F. to l,lO F., in an atmosphere of air, for a period of about 0.5 to about l0 hours.

Although not essential to successful hydroprocessing in all cases, in fact detrimental in some, a halogen component may be incorporated into the catalytic composite. Accordingly, a preferred catalytic composite, for use in the present process, comprises a combination of a Group VIII noble metal component, a germanium component and a halogen component. Although the precise form of the chemistry of the association of the halogen component with the carrier material and metallic components is not accurately known, it is customary in the art to refer to the halogen component as being combined with the carrier material, or with the other ingredients of the catalyst. The combined halogen may be either fluorine, chlorine, iodine, bromine, or mixtures thereof. Of these, fluorine and particularly chlorine are preferred for the hydrocarbon hydroprocesses encompassed by the present invention. The halogen may be added to the carrier material in any suitable manner, and either during preparation of the carrier or before, or after the addition of the other components. For example, the halogen may be added at any stage in the preparation of the carrier material, or to the calcined carrier material, as an aqueous solution of an acid such as hydrogen fluoride, hydrogen chloride, hydrogen bromide, hydrogen iodide, etc. The halogen component or a portion thereof may be composited with the carrier material during the impregnation of the latter with the Group Vlll noble metal component. The inorganic oxide hydrosol, which is typically utilized to form an amorphous carrier material, may contain halogen and thus contribute at least a portion of the halogen component to the final composite. The quantity of halogen is such that the final catalytic composite contains about 0.1 percent to about 1.5 percent by weight, and preferably from about 0.5 percent to about 1.2 percent, calculated on an elemental basis.

With respect to the quantity of the germanium component, it is preferably about 0.01 percent to about 5.0 percent by weight, calculated on an elemental basis. Regardless of the absolute quantities of the germanium component and the platinum group component, the atomic ratio of the Group Vlll noble metal to the germanium contained in the catalyst is preferably selected from the range of about 0. l :l to about 3:1, with excellent results being achieved at an atomic ratio of about 05:1 to a out 1.5:l. This has been found to be particularly true when the total content of the germanium component plus the Group Vlll noble metal component is fixed in the range of about 0.15 to about 3.0 percent by weight. Accordingly, examples of suitable catalytic composites, considering only the Group Vlll noble metal component and the germanium component are as follows: 0.5 percent by weight of germanium, 0.75 percent by weight of platinum; 0.1 percent by weight of germanium, 0.65 percent by weight of platinum; 0.375 percent by weight of germanium, 0.375 percent by weight of platinum; 1.0 percent by weight of germanium, 0.5 percent by weight of platinum; 0.25 percent by weight of germanium, 0.5 percent by weight of platinum; 0.75 percent by weight of palladium, 0.5 percent by weight of germanium; 0.65 percent by weight of palladium, 0.1 percent by weight of germanium; 0.375 percent by weight of palladium, 0.375 percent by weight of germanium; 0.5 percent by weight of palladium, 1.0 percent by weight of germanium; and, 0.5 percent by weight of palladium, 0.25 percent by weight of germanium. When used in many of the hydrogen-consuming processes hereinbefore described, the foregoing quantities of metallic components will be combined with a carrier material of alumina and silica, wherein the silica concentration is l0.0 percent to about 90.0 percent by weight. in those processes wherein the acid function of the catalytic composite must necessarily be attenuated, the metallic components will be combined with a carrier material consisting essentially of alu mina. In this latter situation, a halogen component is often not combined with the catalytic composite, and, the inherent acid function of Group VIII noble metals is further attenuated through the addition of from 0.01 percent to about l.5 percent by weight of an alkalinous metal component.

One such process, in which the acid function of the catalyst employed must necessarily be attenuated, is the process wherein an aromatic hydrocarbon is hydrogenated to produce the corresponding cycloparaffin. Specifically, a benzene concentrate is often used as the starting material for the production of cyclohexane primarily to satisfy the demand therefore in the manufacture of nylon. In order to avoid ring opening which results in loss of both the benzene and the cyclohexane product, an alkalinous metal component is combined with the catalytic composite in an amount of from 0.01 percent to about 1.5 percent by weight. This component is selected from the group of lithium, sodium, potassium, rubidium, cesium, barium, strontium, calcium, magnesium, beryllium, mixtures of two or more, etc. In general, more advantageous results are achieved through the use of the alkali metal, particularly lithium and/or potassium.

In those instances where a halogen component is utilized in the catalyst, it has been determined that more advantageous results are obtained when the halogen content of the catalyst is adjusted during the calcination step through the inclusion of a halogen, or a halogen-containing compound in the air atmosphere. in particular, when the halogen component of the catalyst is chlorine, for example, it is preferred to use a mole ratio of water to hydrochloric acid of about 20:] to about l00:l during at least a portion of the calcination step in order to adjust the final chlorine content of the composite to a range of about 0.5 to about 1.2 percent weight.

Prior to its use, in the hydroprocessing of hydrocarbons, the resultant calcined catalytic composite may be subjected to a substantially water-free reduction technique. This technique is designed to insure a uniform and finely divided dispersion of the metallic components throughout the carrier material. Preferably, substantially pure and dry hydrogen (i.e. less than about 30.0 vol. ppm. of water) is employed as the reducing agent. The calcined catalyst is contacted at a temperature of about 800 F. to about l,000 F., and for a period of time of about 0.5 to about 2 hours, in order to minimize the task of reducing the germanium component, but effected to substantially reduce the Group Vlll noble metal component. This reduction technique may be performed in situ as part of a startup sequence provided precautions are observed to predry the unit to a substantially water-free state.

Again, with respect to effecting hydrogen-consuming reactions, the process is generally improved when the reduced composite is subjected to a presulfiding operation designed to incorporate from about 0.05 to about 0.50 percent by weight of sulfur, on an elemental basis, in the catalytic composite. This presulfiding treatment takes place in the presence of hydrogen and a suitable sulfur-containing compound including hydrogen sulfide, lower molecular weight mercaptans, organic sulfides, etc. The procedure constitutes treating the reduced catalyst with a sulfiding gas, such as a mixture of hydrogen and hydrogen sulfide having about 10 moles of hydrogen per mole of hydrogen sulfide, and at conditions suf ficient to effect the desired incorporation of sulfur. These conditions include a temperature ranging from about 50 F. up to about l,000 F.

According to the present invention, the hydrocarbon charge stock and hydrogen are contacted with a catalyst of the type described above in a hydrocarbon conversion zone. The particular catalyst employed is dependent upon the characteristics of the charge stock as well as the desired end result. The contacting may be accomplished by using the catalyst in a fixed bed system, a moving bed system, a fluidized bed system, or in a batch-type operation; however, in view of the risk of attrition losses of the valuable catalyst, it is preferred to use the fixed bed system. Furthermore, it is well known that a fixed bed catalytic system offers many operational advantages. in this type of system a hydrogen-rich gas and the charge stock are preheated by any suitable heating means to the desired temperature, and then are passed into a conversion zone containing the fixed bed of the catalytic composite. It is understood, of course, that the conversion zone may be one or more separate reactors having suitable means therebetween to insure that the desired conversion temperature is maintained at the entrance to each reactor. It is also to be noted that the reactants may be contacted with the catalyst bed in either upward, downward, or radial flow fashion, with the latter being preferred. Additionally, the reactants may be in the liquid phase, a mixed liquid-vapor phase, or a vapor phase when they contact the catalyst.

The operating conditions imposed upon the reaction zones are dependent upon the particular hydroprocessing being effected. However, these operating conditions will include a pressure from about 400 to about 5,000 p.s.i.g., a liquid hourly space velocity of about 0.1 to about l0.0, and a hydrogen concentration within the range of about 1,000 to about 50,000 standard cubic feet per barrel. In view of the fact that the reactants being effected are exothermic in nature, an increasing temperature gradient is experienced as the hydrogen and feed stock transverses the catalyst bed. For any given hydrogenconsuming process, it is desirable to maintain the maximum catalyst bed temperature below about 900 F., which temperature is virtually identical to that conveniently measured at the outlet of the reaction zone. Hydrogen-consuming processes are conducted at a temperature in the range of about 200 F. to about 900 F., and it is intended herein that the stated temperature of operation alludes to the maximum catalyst bed temperature. In order to assure that the catalyst bed temperature does not exceed the maximum allowed for a given process, the use of conventional quench streams, either normally liquid or gaseous, introduced at one or more intermediate loci of the catalyst bed, may be utilized. in some of the hydrocarbon hydroprocesses encompassed by the present invention, and especially where hydrocracking a heavy hydrocarbonaceous material to produce lower boiling hydrocarbon products, that portion of the normally liquid product effluent boiling above the end point of the desired product will be recycled to combine with the fresh hydrocarbon charge stock. In these situations, the combined liquid feed ration (defined as volumes of total liquid charge to the reaction zone per volume of fresh feed charge to the reaction zone) will be within the range of about [.1 to about 6.0.

Specific operating conditions, processing techniques, particular catalytic composites and other individual process details will be given in the following detailed description of several of the hydrocarbon hydroprocesses to which the present invention is applicable. These will be presented by way of examples given in conjunction with commercially scaled operating units. ln presenting these examples, it is not intended that the invention be limited to the specific illustrations, nor is it intended that a given process be limited to the particular operating conditions, catalytic composite, processing techniques charge stock, etc. It is understood, therefore, that the present invention is merely illustrated by the specifics hereinafter set forth.

EXAMPLEl In this example, the present invention is illustrated as applied to the hydrogenation of aromatic hydrocarbons such as benzene, toluene the various xylenes, naphthalenes, etc., to form the corresponding cyclic paraffins. When applied to the hydrogenation of aromatic hydrocarbons, which are com taminated by sulfurous compounds, primarily thiophenic compounds, the process is advantageous in that it affords 100.0 percent conversion without the necessity for the substantially complete prior removal of the sulfur compoundsv The corresponding cyclic paraffins, resulting from the hydrogenation of the aromatic nuclei, include compounds such as cyclohexane, monodi-, tri-substituted eyclohexanes, decahydronaphthalene, tetrahydronaphthalene, etc., which find widespread use in a variety of commercial industries in the manufacture of nylon, as solvents for various fats, oils, waxes, etc.

Aromatic concentrates are obtained by a multiplicity of techniques. For example, a benzene-containing fraction may be subjected to distillation to provide a heart-cut which contains the benzene, This is then subjected to a solvent extraction process which separates the benzene from the normal or iso-paraffinic components, and the naphthenes contained therein. Benzene is readily recovered from the selected solvent by way of distillation, and in a purity of 99.0 percent or more. Heretofore, the hydrogenation of aromatic hydrocarbons, for example benzene, has been effected with a nickelcontaining catalyst. This is extremely disadvantageous in many respects, and especially from the standpoint that nickel is quite sensitive to the minor quantity of sulfurous compounds which may be contained in the benzene concentrate. In accordance with the present process, the benzene is hydrogenated in contact with a nonacidic catalytic composite containing 0.0l percent to about 2.0 percent by weight of a Group VIII noble metal component, from about 0.01 percent to about 5.0 percent by weight of a germanium component and from about 0.0lpercent to about 1.5 percent by weight of an alkalinous metal component. Operating conditions include a maximum catalyst bed temperature in the range of about 200 F. to about 800 F., a pressure of from 500 to about 2,000 p.s.i.g., a liquid hourly space velocity of about L0 to about 10.0 and a hydrogen concentration in an amount sufficient to yield a mole ratio of hydrogen to cyclohexane, in the product effluent from the last reaction zone, not substantially less than about 4.0:1. Although not essential, one preferred operating technique involves the use of three reaction zones, each of which contains approximately one-third of the total quantity of catalyst employed. The process is further facilitated when the total fresh benzene is added in three approximately equal portions, one each of the inlet of each of the three reaction zones.

The catalyst utilized is a substantially halogen free alumina carrier material combined with 0.50 percent by weight of germanium, 0.375 percent by weight of platinum, and about 0.90 percent by weight oflithium, all of which are calculated on the basis of the elemental metals. The hydrogenation process will be described in connection with a commercially sealed unit having a total fresh benzene feed capacity of about L488 barrels per day. Makeup gas in an amount of about 74 l .6 mols/hr. is admixed with 2,396 Bbl./day (about 329 mols/hr.) of a cyclohexane recycle stream, the mixture being at a temperature of about 137 F., and further mixed with 96.24 mols/hr. (582 BbL/day) of the benzene feed; the final mixture constitutes the total charge to the first reaction zone.

Following suitable heat exchange with various hot effluent streams, the total feed to the first reaction zone is at a temperature of 385 F. and a pressure of 460 p.s.i.g. The reaction zone effluent is at a temperature of 606 F. and a pressure of about 450 p.s.i.g. The total effluent from the first reaction zone is utilized as a heat-exchange medium, in a stream generator, whereby the temperature is reduced to a level of about 545 F. The cooled effluent is admixed with about 98.5 moles per hour (596 Bbl./day) of fresh benzene feed, at a temperature of 100 F.; the resulting temperature is 400 F., and the mixture enters the second reaction zone at a pressure of about 440 psi. g. The second reaction zone effluent, at a pressure of 425 p.s.i.g. and a temperature of 611 F., is admixed with l.2l mols/hr. (310 BbL/day) of fresh benzene feed, the resulting mixture being at a temperature of 578 F. Following its use as a heat exchange medium, the temperature is reduced to 400 F., and the mixture enters the third reaction zone at a pressure of M5 p.s.i.g. The third reaction zone effluent is at a temperature of about 509 F. and a pressure of about 400 p.s.i.g. Through utilization as a heat exchange medium, the temperature is reduced to a level of about 244 F., and subsequently reduced to a level of about ll5 F. by use of aircooled condenser. The cooled third reaction zone effluent is introduced into a high-pressure separator, at a pressure of about 370 p.s.i.g.

A hydrogen-rich vaporous phase is withdrawn from the high pressure separator and recycled by way of compressive means, at a pressure of about 475 p.s.i.g., to the inlet of the first reaction zone. A portion of the normally liquid phase is recycled to the first reaction zone as the cyclohexane concentrate hereinbefore described. The remainder of the normally liquid phase is passed into a stablilizing column functioning at an operating pressure of about 250 p.s.i.g., atop temperature of about 160 F. and a bottom temperature of about 430F. The cyclohexane product is withdrawn from the stabilizer as a bottoms stream, the overhead stream being vented to fuel. The cyclohexane concentrate is recovered in an amount of about 245.80 moles per hour, of which only about 0.60 moles per hour constitutes other hexanes. In brief summation, of the 19,207 pounds per hour of fresh benzene feed, 20,685 pounds per hour of cyclohexane product is recovered.

EXAMPLE ll Another hydrocarbon hydroprocessing scheme, to which the present invention is applicable, involves the hydrorefining of coke-forming hydrocarbon distillates. These hydrocarbon distillates are generally sulfurous in nature, and contain monoolefinic, di-olefinic and aromatic hydrocarbons. Through the utilization of a catalytic composite comprising both a germanium component and a Group Vlll noble metal component, increased selectivity and stability of operation is obtained; selectivity is most noticeable with respect to the retention of aromatics, and in hydrogenating conjugated di-olefinic one mono-olefinic hydrocarbons. Such charge stocks generally result from diverse conversion processes including the catalytic and/or thermal cracking of petroleum, sometimes referred to as pyrolysis, the destructive distillation of wood or coal, shale oil retorting, etc. The impurities in these distillate fractions must necessarily be removed before the distillates are suitable for their intended use, or which when removed, enhance the value of the distillate fraction for further processing. Frequently, it is intended that these charge stocks be substantially desulfurized, saturated to the extent necessary to remove the conjugated di-olefins, while simultaneously retaining the aromatic hydrocarbons. When subjected to hydrorefining for the purpose of removing the contaminating influences, there is encountered difficulty in effecting the desired degree of reaction due to the formation of coke and other carbonaceous material.

As utilized herein, hydrogenating" is intended to be synonymous with hydrorefining. The purpose is to provide a highly selective and stable process for hydrogenating cokeforming hydrocarbon distillates, and this is accomplished through the use of a fixed bed catalytic reaction system utilizing a catalyst comprising a germanium component and a Group Vlll noble metal component. There exists two separate, desirable routes for the treatment of coke-forming distillates, for example a pyrolysis naphtha byproduct. One such route is directed toward a product suitable for use in certain gasoline blending. With this as the desired object, the process can be effected in a single stage, or reaction zone, with the catalytic composite hereinafter specifically described as the first-stage catalyst. The attainable selectivity in this instance resides primarily in the hydrogenation of highly reactive double bonds. In the case of conjugated di-olefins, the selectivity afforded restricts the hydrogenation to produce mono-olefins, and, with respect to the styrenes, for example, the hydrogenation is inhibited to produce alkyl benzenes without ring saturation. The selectivity is accomplished with a minimum of polymer formation either to gums", or lower molecular weight polymers which would necessitate a rerunning of the product before blending to gasoline would be feasible. Other advantages of restricting the hydrogenating of the conjugated di-olefins, such as 1.5 normal hexadiene are not unusually offensive in suitably inhibited gasolincs in some locales, and will not react in this first stage. Some fresh charge stocks are sufficiently low in mercaptan sulfur content that direct gasoline blending may be considered, although a mild treatment for mercaptan sulfur removal might be necessary. These considerations are generally applicable to foreign markets, particularly European, where olefinic and sulfur-containing gasolines are not too objectionable. it must be noted that the sulfurous compounds, and the mono-olefins, whether virgin, or products of di-olefin partial saturation, are unchanged in the single, or first-stage reaction zone. Where however the desired end result is aromatic hydrocarbon retention, intended for subsequent extraction, the two-stage route is required. The mono-olefins must be substantially saturated in the second stage to facilitate aromatic extraction by way of currently utilized methods. Thus, the desired necessary hydrogenation involves saturation of the mono-olefins, as well as sulfur removal, the latter required for an acceptable ultimate aromatic product. Attendant upon this is the necessity to avoid even partial saturation of aromatic nuclei.

With respect to one catalytic composite, its principal function involves the selective hydrogenation of conjugated diolefinic hydrocarbons to mono-olcfinic hydrocarbons. This particular catalytic composite possesses unusual stability notwithstanding the presence of relatively large quantities of sulfurous compounds in the fresh charge stock. The catalytic composite comprises an alumina-containing refractory inorganic oxide, a germanium component, a Group Vlll noble component and an alkali-metal component, the latter being preferably potassium and/or lithium. It is especially preferred, for use in this particular hydrocarbon hydroprocessing scheme, that the catalytic composite be substantially free from any acid-acting propensities. The catalytic composite, util ized in the second reaction zone for the primary purpose of effecting the destructive conversion of sulfurous compounds into hydrogen sulfide and hydrocarbons, is a composite of an alumina-containing refractory inorganic oxide, a Group Vlll noble metal component and a germanium component. Through the utilization of a particular sequence of processing steps, and the use of the foregoing described catalytic composites, the formation of high molecular weight polymers and copolymers is inhibited to a degree which permits processing for an extended period of time. Briefly, this is accomplished by initiating the hydrorefining reactions at temperatures below about 500 F., at which temperatures the coke-forming reactions are not promoted. The operating conditions within the second reaction zone are such that the sulfurous compounds are removed without incurring the detrimental polymerization reactions otherwise resulting were not for the saturation of the conjugated di-olefinic hydrocarbons within the first reaction zone.

The hydrocarbon distillate charge stock, for example a light naphtha byproduct from a commercial cracking unit designed and operated for the production of ethylene, having a gravity of about 340 APl, a bromine number of about 35.0, a diene value of about 17.5 and containing about 1,600 p.p.m. by weight of sulfur and 75.9 vol. percent aromatic hydrocarbons, is admixed with recycled hydrogen. This light naphtha coproduct has an initial boiling point of about 164 F. and an end boiling point of about 333 F. The hydrogen concentration is within the range of from about 1,000 to about 10,000 s.c.f./Bbl., and preferably in the narrower range of from 1,500 to about 6,000 s.c.f./Bbl. The charge stock is heated to a temperature such that the maximum catalyst temperature is in the range of from about 200 F. to about 500 F., by way of heat exchange with various product effluent streams, and is introduced into the first reaction zone at an LHSV in the range of about 0.5 to about 10.0. The reaction zone is maintained at a pressure of from 400 to about 1,000 p.s.i.g., and preferably at a level in the range offrom 500 p.s.i.g. to about 900 p.s.i.g.

The temperature of the product effluent from the first reaction zone is increased to a level above about 500 F., and preferably to result in a maximum catalyst temperature in the range of 600 F. to 900 F. When the process is functioning efficiently, the diene value of the liquid charge entering the second catalytic reaction zone is less than about 1.0 and often less than about 0.3. The conversion of nitrogenous and sulfurous compounds, and the saturation of mono-olefins, contained within the first zone efiluent, is effected in the second zone. The second catalytic reaction zone is maintained under an imposed pressure of from about 400 to about 1,000 p.s.i.g., and preferably at a level of from about 500 to about 900 p.s.i.g. The two-stage process is facilitated when the focal point for pressure control is the high-pressure separator serving to separate the product effluent from the second catalytic reaction zone. It will, therefore, be maintained at a pressure slightly less than the first catalytic reaction zone, as a result of fluid flow through the system. The LHSV through the second reaction zone is about 0.5 to about 10.0, based upon fresh feed only. The hydrogen concentration will be in a range of from 1,000 to about 10,000 s.c.f./Bbl., and preferably from about 1,000 to about 8,000 s.c.f./Bbl. Series-flow through the entire system is facilitated when the recycle hydrogen is admixed with the fresh hydrocarbon charge stock. Makeup hydrogen, to supplant that consumed in the overall process, may be introduced from any suitable external source, but is preferably introduced into the system by way of the effluent line from the first catalytic reaction zone to the second catalytic reaction zone.

With respect to the naphtha-boiling range portion of the product effluent, the sulfur concentration is about 0.1 p.p.m. the aromatic concentration is about 75.1 percent by volume, the bromine number is less than about 0.3 and the diene value is essentially nil."

With charge stocks having exceedingly high diene values, a recycle diluent is employed in order to prevent an excessive temperature rise in the reaction system. Where so utilized, the source of the diluent is preferably a portion of the normally liquid product effluent from the second catalytic reaction zone. The precise quantity of recycle material varies from feed stock to feed stock; however, the rate at any given time is controlled by monitoring the diene value of the combined liquid feed to the first reaction zone. As the diene value exceeds a level of about 25.0, the quantity of recycle is increased, thereby increasing the combined liquid feed ratio; when the diene value approaches a level of about 20.0, or less, the quantity of recycle diluent may be lessened, thereby decreasing the combined liquid feed ratio.

With another so-called pyrolysis gasoline, having a gravity of about 364 APl, containing 600 ppm. by weight of sulfur, 78.5 percent by volume of aromatics, and having a bromine number of 45 and a diene value of 25.5 it is initially processed in a first reaction zone containing a catalytic composite of alumina, 0.5 percent by weight of lithium, 0.2 percent by weight of palladium and 0.375 percent by weight of germanium, calculated as the elements. The fresh feed charge rate is 3,300 Bbl./day, and this is admixed with 2,475 BbL/day of the normally liquid diluent. Based upon fresh feed only, the LHSV is 2.5 and the hydrogen circulation rate is 1,750 s.c.f./Bbl. The charge is raised to a temperature of about 250 F., and enters the first reaction zone at a pressure of about 840 p.s.i.g. The product effluent emanates from the first reaction zone at a pressure of about 830 p.s.i.g. and a temperature of about 350 F. The effluent is admixed with about 660 s.c.f./Bbl. of makeup hydrogen, and the temperature is increased to a level of about 545 F., the heated stream is introduced into the second reaction zone under a pressure of about 790 p.s.i.g. The LHSV, exclusive of the recycle diluent, is 2.5, and the hydrogen circulation rate is about 1,500. The second reaction zone contains a catalyst of a composite of alumina, 0.375 percent by weight of platinum and 0.25 percent by weight of germanium. The reaction product effluent is introduced, following its use as a heat-exchange medium and further cooling, to reduce its temperature from 620 F. to a level of F., into a high-pressure separator at a pressure of about 750 p.s.i.g. The normally liquid stream from the cold separator is introduced into a reboiled stripping column for hydrogen sulfide removal and depentanization. The hydrogen sulfide stripping column functions at conditions of temperature and pressure required to concentrate a C to C aromatic stream as a bottoms fraction. With respect to the overall product distribution, only 690 lbs/hr. of pentanes and lighter hydrocarbons is indicated in the stripper overhead. The aromatic concentrate is recovered in an amount of about 40,070 lbs/hr. (the fresh feed is 40,120 lbs./hr.); these results are achieved with a hydrogen consumption of only 660 s.c.f./Bbl. With respect to the desired product, the aromatic concentration is 78.0 the sulfur concentration is less than 1.0 p.p.m. by weight, and the diene value is essentially nil.

EXAMPLE 111 This example is presented to illustrate still another hydrocarbon hydroprocessing scheme for the improvement of the jet fuel characteristics of sulfurous kerosene boiling range fractions. The improvement is especially noticeable in the IPT Smoke Point, the concentration of aromatic hydrocarbons and the concentration of sulfur. A two-stage process wherein desulfurization is effected in the first reaction zone at relatively mild severities which result in a normally liquid product effluent containing from about 15 to about 35 p.p.m. of sulfur by weight. Aromatic saturation is the principal reaction effected in the second reaction zone, having disposed therein a catalytic composite of alumina, a halogen component, a Group Vlll noble metal component and a germanium component.

Suitable charge stocks are kerosene fractions having an initial boiling point as low as about 300 F., and an end boiling point as high as about 575 F. and, in some instances, up to 600 F. Exemplary of such kerosene fractions are those boiling from about 300 F. to about 550 F., from 330 F. to about 500 F., from 330 F. to about 530 F., etc. As a specific example, a kerosene obtained from hydrocracking a Midcontinent slurry oil, having a gravity of about 305 APl, an initial boiling point of about 388 F., and end boiling point of about 522 F., has an lPT Smoke Point of 17.1 mm., and contains 530 p.p.m. of sulfur and 24.8 percent by volume of aromatic hydrocarbons. Through the use of the catalytic process of the present invention, the improvement in the jet fuel quality of such a kerosene fraction is most significant with respect to raising the IPT Smoke Point, and reducing the concentration of sulfur and the quantity of aromatic hydrocarbons. Specifications regarding the poorest quality of jet fuel, commonly referred to as Jet-A, Jet-A1 and Jet-B call for a sulfur concentration of about 0.3 percent by weight maximum (3,000 p.p.m.) a minimum lPT Smoke Point of 25 mm. and a maximum aromatic concentration of about 20.0 vol. percent.

The charge stock is admixed with recycled hydrogen in an amount within the range of from about l,000 to about 2,000 s.c.f./Bbl. This mixture is heated to a temperature level necessary to control the maximum catalyst bed temperature below about 750 F., and preferably not above 700 F., with a lower catalyst bed temperature of about 600 F. The catalyst, a well known standard desulfurization catalyst containing about 2.2 percent by weight of cobalt and about 5.7 percent by weight of molybdenum, composited with alumina is disposed in a reaction zone maintained under an imposed pressure in the range of from about 500 to about 1,1 p.s.i.g. The LHSV is in the range of about 0.5 to about 10.0, and preferably from about 0.5 to about 5.0. The total product effluent from this first catalytic reaction zone is separated to provide a hydrogen-rich gaseous phase and a normally liquid hydrocarbon stream containing p.p.m. to about 35 p.p.m. of sulfur by weight. The normally liquid phase portion of the first reaction zone ef fluent is utilized as the fresh feed charge stock to the second reaction zone. In this particular instance, the first reaction zone decreases the sulfur concentration to about 25 p.p.m., the aromatic concentration to about 16.3 percent by volume, and has increased the IPT Smoke Point to a level of about 21.5

The catalytic composite within the second reaction zone comprises alumina, 0.375 percent by weight of platinum, 0.30 percent by weight of germanium and about 0.60 percent by weight of combined chloride, calculated on the basis of the elements. The reaction zone is maintained at a pressure of about 400 to about l,500 p.s.i.g., and the hydrogen circulation rate is in the range of 1,500 to about 10.000 s.c.f./Bbl. The LHSV, hereinbefore defined, is in the range of from about 0.5 to about 5 .0, and preferably from about 0.5 to about 3.0. lt is preferred to limit the catalyst bed temperature in the second reaction zone to a maximum level of about 750 F. With a catalyst of this particular chemical and physical characteristics, optimum aromatic saturation, processing a feed stock containing from about l5 to about 35 p.p.m. of sulfur, is effected at maximum catalyst bed temperatures in the range of about 625 F. to about 750 F. With respect to the normally liquid kerosene fraction, recovered from the condensed liquid removed from the total product effluent, the sulfur concentration is effectively nil," being about 0.1 ppm. The quantity of aromatic hydrocarbons has been decreased to a level of about 0.7 percent by volume, and the IPT Smoke Point has been increased to about 26.3 mm.

With respect to another kerosene fraction, having an lPT Smoke Point of about 20.5 mm., an aromatic concentration of about 19.3 vol. percent and a sulfur concentration of about 17 p.p.m. by weight, the same is processed in a catalytic reaction zone at a pressure of about 850 p.s.i.g. and a maximum catalyst bed temperature of about 725 F. The LHSV is about 1.35, and the hydrogen circulation rate is about 8,000 s.c.f./Bbl. The catalytic composite disposed within the reaction zone comprises alumina, 0.25 percent by weight of platinum, 0.40 percent by weight of germanium, about 0.35 percent by weight of combined chloride and 0.35 percent weight of combined fluoride. Following separation and distillation, to concentrate the kerosene fraction, analyses indicate that the Smoke Point has been increased to a level of about 36.9 mm., the aromatic concentration has been lowered to about 0.6 percent by volume and the sulfur concentration is essentially nil.

EXAMPLE IV This illustration of a hydrocarbon hydroprocessing scheme, encompassed by my invention, is one which involves hydrocracking heavy hydrocarbonaceous material into lower boiling hydrocarbon products. ln this instance, the preferred catalysts contain a germanium component, a Group Vlll noble metal component, combined with a crystalline alumina silicate carrier material, preferably faujasite, and still more preferably one which is at least 90.0 percent by weight zeolitic. The

Group Vlll noble metal component is preferably platinum and/or palladium; and, in some instances, a halogen component may be combined therewith, particularly fluorine and/or chlorine.

Most of the virgin stocks, intended for hydrocracking, are contaminated by sulfurous compounds and nitrogenous compounds, and, in the case of the heavier charge stocks, various metallic contaminants, insoluable asphalts, etc. Contaminated charge stocks are generally hydrorefined in order to prepare a charge suitable for hydrocracking. Thus, the catalytic process of the present invention can be beneficially utilized as the second stage of a two-stage process, in the first stage of which the fresh feed is hydrorefined.

Hydrocracking reactions are generally effected at elevated pressures in the range of about 800 to about 5,000 p.s.i.g., and preferably at some intermediate level of 1,000 to about 3,500 p.s.i.g. Liquid hourly space velocities of about 0.25 to about 10.0 will be suitable, the lower range generally reserved for the heavier stocks. The hydrogen circulation rate will be at least about 3,000 s.c.f./Bbl. with an upper limit of about 50,000 s.c.f./Bbl., based upon fresh feed. For the majority of feed stocks, hydrogen concentrations in the range of 5,000 to 20,000 s.c.f./Bbl. will suffice. With respect to the LHSV, it is based upon fresh feed, notwithstanding the use of recycle liquid providing a combined liquid feed ratio in the range of about 1.25 to about 6.0. The operating temperature again alludes to the temperature of the catalyst within the reaction zone, and is in the range of about 400 F. to about 900 F. Since the principal reactions are exothermic in nature, the increasing temperature gradient, experienced as the charge stock traverses the catalyst bed, results in an outlet temperature higher than that at the inlet to the catalyst bed. The maximum catalyst temperature should not exceed 900 F., and it is generally a preferred technique to limit the temperature increase to F. or less.

Although amorphous composites of alumina and silica, containing from about 10.0 percent to about 90.0 percent by weight of the latter, are suitable for use in the catalytic composite employed in the present process, a preferred carrier material constitutes a crystalline aluminosilicatc, preferably faujasite, of which at least about 90.0 percent by weight is zeolitic. This carrier material, and a method of preparing the same, have hereinbefore described. Generally, the germanium component will be used in an amount sufficient to result result in a final catalytic composite containing about 0.0l percent to about 5.0 percent by weight. The Group Vlll noble metal component is generally present in an amount within the range of about 0.01 percent to about 2.0 percent by weight, and may exist within the composite as a compound such as an oxide, sulfide, halide, etc. Another possible constituent of the catalyst is a halogen component, either fluorine, chlorine, iodine, bromine, or mixtures thereof. Of these it is preferred to utilize a catalyst containing fluorine and/or chlorine. The halogen component will be composited with the carrier material in such a manner as results in a final composite containing about 0.1 percent to about 1.5 percent by weight of halogen, calculated on an elemental basis.

A specific illustration of this hydrocarbon hydroprocessing technique involves the use of a catalytic composite of about 0.4 percent by weight of platinum, 0.7 percent by weight of combine chlorine, and 0.4 percent by weight of germanium, combined with a crystalline aluminosilicate material of which about 90.9 percent by weight constitutes faujasite. This catalyst is intended for utilization in the conversion of l6,000 BbL/day ofa blend oflight gas oils to produce maximum quantities of a heptane-400 F. gasoline-boiling range fraction. The charge stock has a gravity of 338 AH, contains 0.l9 percent by weight of sulfur (1,900 p.p.m.) and 67 p.p.m. by weight of nitrogen, and has an initial boiling point of 369 F., a 50 percent volumetric distillation temperature of 494 F. and an end boiling point of 655 F. The charge stock is initially subjected to a clean-up operation at maximum catalyst temperature of 750 F., a combined feed ratio of 1.0 an LHSV of 2.41 with a hydrogen circulation rate of about 5,000 s.c.f./Bbl. The pressure imposed upon the catalyst within the reaction zone is about 1,500 p.s.i.g. Since at least a portion of the blended gas oil charge stock will be converted into lower boiling hydrocarbon products, the effluent from this cleanup reaction zone is separated to provide a normally liquid, 400 F. plus charge for the hydrocracking reaction zone containing the platinum-germanium-chloride catalyst. The pressure imposed upon the second reaction zone is about 1,500 p.s.i.g. and the hydrogen circulation rate is about 8,000 s.c.f./Bbl. The original quantity of fresh feed to the cleanup reaction zone is about 16,000 BbL/day; following separation of the first zone effluent to provide the 400 F. plus charge to the second reaction zone, the charge to the second reaction zone is in an amount of about 12,150 BbL/day, providing an LHSV of 0.85. The temperature at the inlet to the catalyst bed is 665 F., and a conventional hydrogen quench stream is utilized to maintain the maximum reactor outlet temperature at about 700 F. Following separation of the product effluent from the second reaction zone, to concentrate the desired gasoline boiling range fraction, the remaining 400 F. plus normally liquid material, in an amount of 7,290 BbL/day, is recycled to the inlet of the second reaction zone, thus providing a combined liquid feed ratio thereto of about 1.60. In the following table, there is indicated the product yield and distribution of this process. With respect to normally liquid hydrocarbons, for convenience including butanes, the yields are given in volume percent; with respect to the normally gaseous hydrocarbons, ammonia and hydrogen sulfide, the yields are given in terms of weight percent. With respect to the first reaction zone, the hydrogen consumption is 1.31 percent by weight of the fresh feed (741 s.c.f./Bbl.), and for the hydrocracking reaction zone, 1.26 percent by weight of the fresh feed charge stock, or 713 s.c.f./Bbl.

TABLE: l-lydrocraeking Product Yield and Distribution Charge to Stage ll With respect to both the butane product and pentane product, the former is indicated as being about 68.0 percent isobutanes, while the latter constitutes about 93.0 percent isopentanes. An analysis of the combined pentane/hexane fraction indicates a gravity of 82.6 APl, a clear research octane rating of 85.0 and leaded research octane rating of 99.0; it will be noted that this constitutes an excellent blending component for motor fuel. The desired heptane-400 F. product indicates a gravity of 488 APl, a clear research octane rating of 72.0 and a leaded research octane rating of 88.0. This gasoline boiling range fraction constitutes about 34.0 percent by volume paraffins, 36.0 percent volume napthenes and 30.0 percent by volume aromatic hydrocarbons. It will be recognized that this gasoline-boiling range fraction constitutes an excellent charge stock for a catalytic reforming unit to improve the mot or fuel characteristics thereof.

The foregoing specification, and particularly the examples,

indicates the method by which the present invention is ef fected, and the benefits afforded through the utilization thereof.

1 claim as my invention:

l. A process for hydrocracking a hydrocarbonaceous charge stock into lower molecular weight hydrocarbons which comprises reacting said charge stock with hydrogen, at a temperature of about 400to 900 F., a pressure of about 800 to 5,000 p.s.i.g., a liquid hourly space velocity of about 0.25 to 10.0 and a hydrogen concentration of about 3,000 to 50,000 s.c.f./Bbl., in contact with a catalytic composite of from 0.01 to about 2.0 wt. percent of a group Vlll noble metal component, from 0.01 to about 5.0 wt. percent of a germanium component, and a porous carrier material, said weight percentages being on an elemental basis.

2. The process of claim 1 further characterized in that said Group VIII noble metal is platinum or palladium.

3. The process of claim 2 further characterized in that said catalytic composite is reduced and sulfided prior to contacting said charge stock.

4. The process of claim 1 further characterized in that said catalytic composite contains from about 0.1 percent to about 1.5 percent by weight of a halogen component, on an elemen' tal basis.

5. The process of claim 2 further characterized in that said carrier material is crystalline aluminosilicate.

6. The process of claim 2 further characterized in that said carrier material is an amorphous refractory inorganic oxide.

7. A catalytic composite comprising a substantially pure crystalline aluminosilicate carrier material, at least about 90.0 percent by weight of which is zeolitic, about 0.01 to about 2.0 wt. percent of a Group Vlll noble metal component, and about 0.01 to about 5.0 wt. percent of a germanium component, on an elemental basis.

8. The catalytic composite of claim 7 further characterized in that said Group Vll noble metal component is a platinum component.

9. The catalytic composite of claim 7 further characterized in that said crystalline aluminosilicate is faujasite. 

2. The process of claim 1 further characterized in that said Group VIII noble metal is platinum or palladium.
 3. The process of claim 2 further characterized in that said catalytic composite is reduced and sulfided prior to contacting said charge stock.
 4. The process of claim 1 further characterized in that said catalytic composite contains from about 0.1 percent to about 1.5 percent by weight of a halogen component, on an elemental basis.
 5. The process of claim 2 further characterized in that said carrier material is crystalline aluminosilicate.
 6. The process of claim 2 further characterized in that said carrier material is an amorphous refractory inorganic oxide.
 7. A catalytic composite comprising a substantially pure crystalline aluminosilicate carrier material, at least about 90.0 percent by weight of which is zeolitic, about 0.01 to about 2.0 wt. percent of a Group VIII noble metal component, and about 0.01 to about 5.0 wt. percent of a germanium component, on an elemental basis.
 8. The catalytic composite of claim 7 further characterized in that said Group VII noble metal component is a platinum component.
 9. The catalytic composite of claim 7 further characterized in that said crystalline aluminosilicate is faujasite. 